Process for increasing hydrogen content of synthesis gas

ABSTRACT

Process for increasing the hydrogen content of a synthesis gas containing one or more sulphur compounds, the synthesis gas including hydrogen, carbon oxides and steam, and having a ratio defined as R═(H 2 —C0 2 )/(CO+C0 2 )≦ 0.6  and a steam to carbon monoxide ratio ≦1.8, includes the steps of (i) adjusting the temperature of the synthesis gas; (ii) passing at least a portion of the heated synthesis gas adiabatically through a first bed of sulphur-tolerant water-gas shift catalyst disposed in a first shift vessel at a space velocity ≧12,500 hour −1  to form a pre-shifted gas stream; and (iii) forming a shifted gas stream by subjecting at least a portion of the pre-shifted gas stream to a second stage of water-gas shift in a second shift vessel containing a second bed of sulphur-tolerant water-gas shift catalyst that is cooled in heat exchange with a gas stream including the synthesis gas.

PROCESS

This invention relates to a process for increasing the hydrogen contentof a synthesis gas, in particular increasing the hydrogen content of asynthesis gas generated from a carbonaceous feedstock.

Synthesis gas, also termed syngas, comprising hydrogen and carbon oxides(CO and CO₂) may be generated by a gasification of carbonaceousfeedstocks such as coal, petroleum coke or other carbon-rich feedstocksusing oxygen or air and steam at elevated temperature and pressure.Generally, the resulting synthesis gas is hydrogen deficient and toincrease the concentration of hydrogen, it is necessary to subject theraw synthesis gas to the water-gas-shift reaction by passing it, in thepresence of steam, over a suitable water-gas shift catalyst at elevatedtemperature and pressure. The CO₂ that is formed may then be removed ina downstream gas washing unit to give a hydrogen rich product gas. Thesynthesis gas generally contains one or more sulphur compounds and somust be processed using sulphur-tolerant catalysts, known as “sourshift” catalysts. The reaction may be depicted as follows;

H₂O+CO

H₂+CO₂

This reaction is exothermic, and conventionally it has been allowed torun adiabatically, with control of the exit temperature governed by feedgas inlet temperature and composition.

Furthermore, where it is required that only fractional shift conversionis needed to achieve a target gas composition, this is conventionallyachieved by by-passing some of the synthesis gas around the reactor.

Side reactions can occur, particularly methanation, which is usuallyundesirable. To avoid this, the shift reaction requires considerableamounts of steam to be added to ensure the desired synthesis gascomposition is obtained with minimum formation of additional methane.The cost of generating steam can be considerable and therefore there isa desire to reduce the steam addition where possible.

WO2010/106148 discloses a process to prepare a hydrogen rich gas mixturefrom a halogen containing gas mixture comprising hydrogen and at least50 vol. % carbon monoxide, on a dry basis, by contacting the halogencontaining gas mixture with water having a temperature of between 150and 250 DEG C. to obtain a gas mixture poor in halogen and having asteam to carbon monoxide molar ratio of between 0.2:1 and 0.9:1 andsubjecting said gas mixture poor in halogen to a water-gas shiftreaction wherein part or all of the carbon monoxide is converted withthe steam to hydrogen and carbon dioxide in the presence of a catalystas present in one fixed bed reactor or in a series of more than onefixed bed reactors and wherein the temperature of the gas mixture as itenters the reactor or reactors is between 190 and 230 degrees C. Thespace velocity in the water-gas shift reactor is preferably between6000-9000 h⁻¹. In the single Example a space velocity of 8000 hr⁻¹ wasused. Because this process operates at a low steam to CO ratio and atlow inlet temperature it is limited in utility to certain types ofgasifier and requires a relatively high catalyst volume. Therefore thereis a need for a process operating at a low steam to CO ratio thatrequires less catalyst and which has broader utility.

WO2010/013026 discloses a process for increasing the hydrogen content ofa synthesis gas containing one or more sulphur compounds, comprising thesteps of (i) heating the synthesis gas and (ii) passing at least part ofthe heated synthesis gas and steam through a reactor containing a sourshift catalyst, wherein the synthesis gas is heated by passing itthrough a plurality of tubes disposed within said catalyst in adirection co-current to the flow of said synthesis gas through thecatalyst. The resulting synthesis gas may be passed to one or moreadditional reactors containing sour shift catalyst to maximise the yieldof hydrogen production, or used for methanol production, for theFischer-Tropsch synthesis of liquid hydrocarbons or for the productionof synthetic natural gas. While effective, we have found that in somecases with a cooled first shift reactor that the catalyst temperatureprofile may be too high, leading to undesirable side-reactions.

We have found that the disadvantages of the previous processes may beovercome using a pre-shift stage operated at a high gas hourly spacevelocity in combination with a downstream gas-cooled shift vessel.

Accordingly, the invention provides a process for increasing thehydrogen content of a synthesis gas containing one or more sulphurcompounds, said synthesis gas comprising hydrogen, carbon oxides andsteam, and having a ratio, R, defined as R═(H₂—CO₂)/(CO+CO₂)≦0.6 and asteam to carbon monoxide ratio ≦1.8, comprising the steps of (i) heatingthe synthesis gas, (ii) passing at least a portion of the heatedsynthesis gas adiabatically through a first bed of sulphur-tolerantwater-gas shift catalyst disposed in a first shift vessel at a spacevelocity ≧12,500 hour⁻¹ to form a pre-shifted gas stream, and (iii)forming a shifted gas stream by subjecting at least a portion of thepre-shifted gas stream to a second stage of water-gas shift in a secondshift vessel containing a second bed of sulphur-tolerant water-gas shiftcatalyst that is cooled in heat exchange with a gas stream comprisingthe synthesis gas.

In the present invention the synthesis gas comprising hydrogen andcarbon oxides and containing one or more sulphur compounds may beproduced by any method although it is particularly suited to synthesisgas produced by gasification of a carbonaceous feedstock at elevatedtemperature and pressure. Any known gasification technology may be used.The carbonaceous feedstock may be coal, petroleum coke or anothercarbon-rich feedstock.

Preferably the carbonaceous feedstock is a coal. In coal gasification, acoal powder or aqueous slurry may be partially combusted in a gasifierin a non-catalytic process using oxygen or air and in the presence ofsteam at pressures up to about 85 bar abs and exit temperatures up toabout 1450° C., preferably up to about 1400° C., to generate a rawsynthesis gas comprising hydrogen and carbon oxides (carbon monoxide andcarbon dioxide) and containing one or more sulphur compounds such ashydrogen sulphide and carbonyl sulphide.

The R ratio, defined as R═(H₂—CO₂)/(CO+CO₂), in the synthesis gas feedis ≦0.6 and is preferably in the range 0.1 to 0.6, more preferably 0.2to 0.6. R may readily be calculated from the molar quantities of thecomponents in the synthesis gas feed.

Before the synthesis gas is subjected to the water-gas shift reaction,it is preferably cooled, optionally filtered, then washed to removeparticulates such as coal ash.

The synthesis gas comprises one or more sulphur compounds, such ashydrogen sulphide. In order that the water-gas shift catalysts remainsuitably sulphided, the sulphur content of the synthesis gas fed to thewater-gas shift catalyst is desirably >250 ppm.

If the synthesis gas does not contain enough steam for the water-gasshift process, steam may be added to the synthesis gas, for example bylive steam addition or saturation or a combination of these. Steam maybe added to the synthesis gas before or after heating in the secondvessel. The steam to carbon monoxide ratio (i.e. molar ratio) of thesynthesis gas mixture fed to the first water-gas shift catalyst shouldbe ≦1.8 and preferably is in the range 0.2 to 1.8, more preferably 0.7to 1.8. In some embodiments, it may be desirable to operate with a ratioin the range 0.95 to 1.8.

The water-gas shift catalyst used in the shift vessels may be anysuitably stable and active water-gas shift catalyst. The synthesis gascontains one or more sulphur compounds and so the water-gas shiftcatalyst should remain effective in the presence of these compounds. Inparticular so-called “sour shift” catalysts may be used, in which theactive components are metal sulphides. Preferably the water-gas shiftcatalyst comprises a supported cobalt-molybdenum catalyst that formsmolybdenum sulphide in-situ by reaction with hydrogen sulphide presentin the synthesis gas stream. The Co content is preferably 2-8% wt andthe Mo content preferably 5-20% wt. Alkali metal promoters may also bepresent at 1-10% wt. Suitable supports comprise one or more of alumina,magnesia, magnesium aluminate spinel and titania. The catalysts may besupplied in oxidic form, in which case they require a sulphiding step,or they may be supplied in a pre-sulphided form. Particularly preferredsour shift catalysts are supported cobalt-molybdate catalysts such asKATALCO™ K8-11 available from Johnson Matthey PLC, which comprises about3% wt. CoO and about 10% wt. MoO₃ supported on a particulate supportcontaining magnesia and alumina.

It is desirable to adjust the temperature of the synthesis gas so thatthe temperature within the first water-gas shift vessel is maintainedwithin suitable operating conditions. For instance, after the synthesisgas is washed, thereby significantly cooling it, it may be advantageousto preheat the synthesis gas passing to the vessel. A suitable heatexchanger can be placed on the feed synthesis gas stream. According tothe particular details of the process, suitable media for heat exchangewith the inlet gas may be, for example, another gas stream at adifferent temperature, steam or water. Furthermore, using such a heatexchanger, with a bypass provided around it, gives the ability tocontrol the inlet temperature to the catalyst bed, independently ofvariation in other parameters.

In the present invention, at least part of the temperature adjustment ofthe synthesis gas before it is fed to the first shift vessel includesheating it by passing the synthesis gas through heat exchange apparatus,such as a plurality of tubes, coils or plates, disposed within thesecond catalyst bed. The synthesis gas is at a lower temperature thanthe reacting pre-shifted gas stream and accordingly the synthesis gasacts as a cooling medium thereby removing heat from the second catalystbed. A preferred temperature for the synthesis gas fed to the heatexchange apparatus within the second catalyst bed is in the range 150 to250° C. The synthesis gas is heated as it passes through the heatexchange apparatus in the second vessel. The heated synthesis gasrecovered from the heat exchange apparatus in the second vessel may befurther heated or cooled to provide the desired inlet temperature forthe first shift vessel.

The inlet temperature for the first bed of water-gas shift catalyst maybe in the range 190 to 350° C., preferably 200 to 330° C.

If desired, the heated synthesis gas recovered from the heat exchangeapparatus in the second vessel may be divided into first and secondstreams, with the first stream passed over the first bed of shiftcatalyst and the second stream by-passing the first bed of shiftcatalyst, thereby forming a catalyst by-pass stream. In addition oralternatively, it may be desirable, upstream of the catalyst of thesecond shift vessel, to divide the synthesis gas into first and secondstreams, with the first stream fed to the second shift vessel where itis heated, and the second stream by-passing the second shift vessel,thereby forming a vessel by-pass stream. The catalyst by-pass stream mayif desired be combined with the vessel by-pass stream, thereby forming acombined by-pass stream. The combined by-pass stream is preferably ≦40%by volume of the total synthesis gas feed.

The by-pass stream may be fed to one or more of the pre-shifted gasstream, the shifted gas stream, or separately to downstream processes.Utilising a vessel by-pass around the second shift stage or a combinedby-pass stream around both first and second stage shift vessels isuseful when it is desired to precisely control the overall extent of COconversion for the total synthesis gas feed, e.g. for making synthesisgas of a specific R ratio, as for methanol synthesis. Especially usefulis a use of a second vessel by-pass, because this also allows bettercontrol of the gas flow, so that the temperature profile in thepre-shift vessel is unaffected by control of the extent of COconversion.

If desired, a by-pass stream may be subjected to a carbonyl sulphide(COS) hydrolysis step by passing the stream over a COS hydrolysiscatalyst, such as a particulate alumina or titania based catalyst,disposed in a suitable vessel. In this step, the COS in the by-passstream is hydrolysed by steam to form H₂S, which may be easier to removein downstream processes. In such a COS hydrolysis step, essentially nowater-gas shift reaction takes place.

The synthesis gas and steam mixture is passed at elevated temperatureand pressure, preferably temperatures in the range 190 to 420° C. morepreferably 200 to 400° C., and pressure up to about 85 bar abs, over thefirst bed of water-gas shift catalyst. The flow-rate of synthesis gascontaining steam should be such that the gas hourly space velocity(GHSV) is ≧12,500 hour⁻¹, and is preferably ≧15,500 hour⁻¹, morepreferably ≧17,500 hour⁻¹, most preferably ≧20,000 hour⁻¹.

The water-gas shift reaction occurs, consuming carbon monoxide and steamand forming carbon dioxide and hydrogen. Under the conditions, only aportion of the carbon monoxide and steam are consumed and so thepre-shifted gas stream comprises hydrogen, carbon monoxide, carbondioxide and steam that may be further reacted in the one or more furtherstages of water-gas shift. Under the reaction conditions it is desirableto convert 10 to 40% (by moles) of the carbon monoxide present in thesynthesis gas to carbon dioxide over the first bed of water-gas shiftcatalyst disposed in the first shift vessel. The first shift vessel maythus be termed a pre-shift vessel.

The pre-shift vessel operates adiabatically without applied cooling andso the reacting gases are heated as they pass through the first shiftvessel. Thus some cooling of the pre-shifted gas may be desirable beforepassing the pre-shifted gas stream to the second stage of water-gasshift.

At least a portion of the pre-shifted synthesis gas from the reactorcontaining the first bed of sulphur-tolerant water-gas shift catalyst isfed to the second water-gas shift stage. In the second stage thepre-shifted gas stream is further reacted over a second bed ofsulphur-tolerant water-gas shift catalyst. The second bed of catalystmay be the same or different to the first bed but is preferably also asupported cobalt-molybdenum water-gas shift catalyst. The bed is cooledby heat exchange with a gas stream comprising the synthesis gas fed tothe first shift vessel. The gas stream comprising the synthesis gas maybe passed through tubes, coils or plates. Axial or radial flow vesselsmay be used. Axial flow vessels comprising a plurality of vertical tubesthough which the gas stream comprising synthesis gas flows arepreferred. Where tubes are used, the synthesis gas may be passed throughthe tubes in the second shift vessel in a direction that iscounter-current or co-current to the flow of the pre-shifted gas streamthrough the vessel. Axial, co-current flow is preferred. Where tubes areused, heat transfer enhancement devices may be used inside the tubes,for example core-rods or structures which increase the turbulence of theflowing gas within the tubes

If desired, additional steam may be added to the pre-shifted gas streambefore the second stage of water-gas shift.

The second shift vessel is preferably operated at a temperature in therange 250 to 420° C., more preferably 340 to 400° C. The gas-hourlyspace velocity in the second bed of sulphur-tolerant water-gas shiftcatalyst may be ≧5000 h⁻¹, preferably ≧6000 h⁻¹ and is more preferablyin the range 6000 h⁻¹ to 12000 h⁻¹, most preferably 6000 h⁻¹ to 10000h⁻¹.

The resulting shifted gas stream from the second water-gas shift vesselmay be used in downstream processes for the production of methanol,dimethylether (DME), Fischer-Tropsch (FT) liquids or synthetic naturalgas (SNG). Where a higher degree of water-gas shift is required, forexample when making hydrogen or a low carbon content fuel for combustionin a gas turbine, additional water-gas shift steps may be performed. Insuch cases, one or more further water-gas shift stages, which may beuncooled or cooled and operated in series or parallel, may be used.Preferably one or two further stages of adiabatic water-gas shift areused in series, with optional cooling before each stage, to maximise COconversion in the shifted gas stream.

The present invention has a number of distinct advantages over the priorart processes. Heat generation in each of the first two shift stages isless and therefore it is easier to control the peak temperature in eachbed, and thus minimise the formation of by-products. As a result ofusing the cooled second shift vessel, a more optimal water-gas shiftreaction profile is followed without the use of excessive heat transferarea. This permits a reduction in the total volume of catalyst or agreater CO conversion achieved with the same catalyst volume. Vesseland/or tube inlet temperatures can be varied over a wide range in orderto accommodate varying catalyst reaction activity, without the risk ofhigh peak temperatures.

In the present invention, the feed synthesis gas is preheated, whilecooling the catalyst in the second shift reactor. This is in contrast toa process scheme with sequential adiabatic reactors, where the wholefeed gas stream must be heated in a separate heat exchanger to the inlettemperature to the first shift reactor. Hence, the use of a Gas CooledReactor in this invention reduces the equipment count by combining thepreheating duty within a shift reactor.

The process of the present invention does not rely on having a very lowH2O/CO ratio in the feed gas to limit the theoretical equilibrium COconversion and associated temperature rise. It is also applicable to awide range of gasifier types, including those with a radiant cooling andquench section, which therefore have a higher, involuntary water contentand are unsuitable for utilising the ‘steam deficient’ shift methodologyset out in the aforesaid WO2010/106148. In order to generate ahydrogen-rich synthesis gas the process preferably further comprises thesteps of:

-   -   (i) cooling a shifted gas stream or a mixture of the shifted gas        stream and a bypass stream, to below the dew point to condense        water,    -   (ii) separating the resulting condensate therefrom to form a dry        gas stream,    -   (iii) feeding the dry gas stream to a gas-washing unit operating        by means of counter-current solvent flow, to produce a product        synthesis gas and    -   (iv) collecting the product synthesis gas from the washing unit.

The shifted gas stream may be subjected to these steps alone to form adry shifted gas stream, or as a mixture with a bypass stream.Alternatively, a bypass stream may be separately subjected to thesesteps to form a dry un-shifted by-pass stream, which is fed to the sameor a separate gas washing unit. Where the dry un-shifted gas is fed tothe same gas washing unit, preferably this un-shifted stream is fed tothe gas washing unit such that the solvent flowing through said unitcontacts first with the dry un-shifted synthesis gas and then the dryshifted gas stream.

The cooling step may be performed by heat exchange, e.g. with coldwater, to cool the gases to below the dew point at which steamcondenses. The resulting condensates, which comprise water and somecontaminants, are separated.

The gases may be further cooled and dried, e.g. by means of chilledsolvent, and then fed to a gas-washing unit operating by means ofcounter-current solvent flow. In the gas-washing unit, also known as anacid-gas removal (AGR) unit, a solvent suitable for thedissolution/absorption of carbon dioxide flows counter-current to gasflowing through the unit and dissolves/absorbs carbon dioxide present inthe gas stream. A small quantity of other gas components in the gasstream, particularly carbon monoxide, will also be co-absorbed.Contaminants present in the gas stream that may poison downstreamcatalysts, e.g. sulphur compounds such as H₂S & COS, may also be removedto differing extents. Using AGR, CO₂ levels may be reduced to below 5mole%, on a dry gas basis.

Suitable solvents for absorbing CO₂ are physical solvents, includingmethanol, other alcohol or glycol products, such as glycols orpolyethylene glycol ethers, and propylene carbonate, and chemicalsolvents, such as activated alkanolamines. Methanol is the preferredsolvent where a downstream catalyst is being used. Methanol may be usedat temperatures in the range −30 to −70° C. and at elevated pressures upto about 75 bar abs.

A gas-washing unit may comprise, for example, a column having a solventinlet near the top and a solvent outlet near the bottom, down which asolvent suitable for the dissolution/absorption of carbon dioxide flowsover one or more perforate trays or packing. The gases passing upthrough the column contact the solvent and carbon dioxide isdissolved/absorbed. The gases may leave the column near the top via asynthesis gas outlet. The synthesis gas is cold and may be used to coolthe feed gases to the gas-washing unit using suitable heat exchangemeans such as a spiral wound heat exchanger. In one embodiment, the dryby-pass synthesis gas mixture and dry shifted gas stream are fedseparately to the unit, with the separate feeds arranged such that thatthe solvent contacts first with the dry by-pass synthesis gas mixtureand then the dry shifted gas stream. This is in contrast to previousprocesses, where a synthesis gas mixture is fed to a gas-washing unit sothat the solvent contacts the gas mixture in one stage. We have foundthat by separately feeding the two different gas streams to the unitsuch that that the solvent contacts first with the dry gas mixture andthen the dry shifted gas stream, the efficiency of the process isimproved, which offers the potential for reduced CO co-absorption and anincreased potential for methanol or liquid hydrocarbon production from agiven quantity of synthesis gas.

The process is desirably operated such that the synthesis gas collectedfrom the gas-washing unit has an R ratio suited to the downstream use,such as methanol or DME production, FT hydrocarbon production or SNGproduction. For the production of methanol or hydrocarbons, the desiredstoichiometry ratio, R, of the product synthesis gas is preferably inthe range 1.4 to 2.5. For generating synthetic natural gas (SNG) therange is preferably in the range 2.8 to 3.3. Alternatively, the sourshift reactor, additional downstream sour shift stage or stages, andgas-washing stage may be operated such that the synthesis gas collectedfrom the gas-washing unit is hydrogen rich, with minimal CO and CO₂content, where this is desirable. Such hydrogen-rich gas streams may beused in ammonia synthesis, for hydrogenation purposes, for chemicalssynthesis or power generation by combustion in a gas turbine with orwithout additional hydrocarbon fuels.

The invention is further illustrated by reference to the accompanyingdrawings in which;

FIG. 1 is a depiction of one embodiment according to the presentinvention suitable for feed from a gasifier having a radiant cooling andquench section producing a steam-containing synthesis gas with asteam:CO ratio in the range 1.3-1.4 and operating with co-current flowthrough tubes disposed within the second vessel, and

FIG. 2 is a depiction of a further embodiment suitable for feed fromgasifier producing a steam-containing synthesis gas with a steam:COratio in the range 0.20-0.30 and operating with counter-current flowthrough tubes disposed the second vessel.

In FIG. 1, a synthesis gas 110 containing one or more sulphur compoundsand steam with a steam:CO ratio in the range 1.3-1.4 is fed to adistributor 112 disposed within a second sour shift vessel 114. Thedistributor is connected to a plurality of tubes 116 that passvertically through a bed of particulate Co/Mo sour shift catalyst 118.The synthesis gas is able to pass from the distributor verticallythrough the tubes where it is heated thereby cooling the reactant gas inthe catalyst bed 118. The tubes are connected to a collector 120 at theother end of the tubes that collects heated synthesis gas.

The heated synthesis gas is fed via line 122 to heat exchanger 124 whereits temperature is adjusted to the desired inlet temperature for thefirst water-gas shift catalyst. The temperature adjusted synthesis gasis fed from exchanger 124 via line 126 to a pre-shift vessel 128containing a first fixed bed of particulate sulphided Co/Mosulphur-tolerant water-gas shift catalyst 130. The flow of synthesis gascontaining steam is controlled such that the gas hourly space velocityin the first bed of catalyst is >12,500 h⁻¹. The synthesis gascontaining steam reacts over the catalyst to form carbon dioxide andhydrogen. The pre-shift vessel is operated adiabatically without coolingand the exothermic reactions heat the resulting pre-shifted gas stream.

The pre-shifted gas stream is recovered from the vessel 128 via line 132and passed through heat exchanger 134 where it is cooled.

The cooled pre-shifted gas stream is then fed via line 136 to the inletof the second water-gas shift vessel 114 containing the second fixed bedof particulate sulphided Co/Mo sulphur-tolerant water-gas shift catalyst118. If desired, additional steam may be added to the pre-shifted gasmixture 132 upstream of vessel 114 (not shown). The pre-shifted gasmixture is passed over the water-gas shift catalyst 118 furtherincreasing the hydrogen content of the synthesis gas. The bed ofcatalyst 118 is cooled in heat exchange with the synthesis gas 110passing through the tubes 116 in a direction co-current to the flow ofpre-shifted gas stream through the vessel 114. A hydrogen-enrichedshifted gas stream is recovered from the outlet of the second vessel 114via line 138.

The shifted gas stream is cooled in heat exchangers 140 and 142 and thecooled shifted synthesis gas stream fed via line 144 to a thirdwater-gas shift vessel 146 containing a third particulate bed ofsulphur-tolerant Co/Mo water-gas shift catalyst 148. The shifted gasstream containing steam further reacts over the catalyst 148 to formcarbon dioxide and hydrogen. The third vessel is operated adiabaticallywithout cooling and the exothermic reactions heat the resulting shiftedgas stream. The shifted gas stream is recovered from the third water-gasshift vessel 146 and passed via line 150 to heat exchanger 152 where itis cooled. The cooled shifted gas stream is then fed via line 154 to afourth water-gas shift vessel 156 containing a fourth particulate bed ofsulphur-tolerant Co/Mo water-gas shift catalyst 158. The shifted gasstream containing steam further reacts over the catalyst 158 to formcarbon dioxide and hydrogen. The fourth vessel is operated adiabaticallywithout cooling and the exothermic reactions heat the resulting shiftedgas stream.

The shifted gas stream is recovered from the fourth water-gas shiftvessel 156 via line 160 and passed through heat exchanger 162, andoptionally further heat exchangers (not shown) to cool the gas below thedew point and so condense the remaining steam. The cooled shifted streamis fed via line 164 to separator 166 in which the condensate isseparated from the hydrogen rich shifted gas stream. The dryhydrogen-rich shifted gas stream is recovered from separator 166 vialine 168 and the condensate via line 170. The condensate may be used togenerate steam for use in the process. The dry hydrogen-rich shifted gasstream 168 may be used in downstream processing or sent to a gas washingunit (not shown) to recover CO2 and H2S and generate a hydrogen rich gasstream product. The carbon dioxide recovered from such processes may beused in carbon-capture and storage (CCS) processes or in enhanced oilrecovery (EOR) processes.

In an alternative embodiment, by utilising the collector 120 as thedistributor and vice-versa, the synthesis gas 110 may be fed through thetubes 116 in a direction counter-current to the flow of pre-shifted gasthrough the second water-gas shift vessel 114.

In FIG. 2 the process is modified by steam addition to the synthesis gasbefore and after heating in the second shift vessel. Accordingly, asynthesis gas 210 containing one or more sulphur compounds with asteam:CO ratio in the range 0.20-0.30 is heated in heat exchanger 212and mixed with steam from line 214. The steam in line 214 is provided bya boiler-feed water supply 216 heated by heat exchanger 218. Additionalsteam is supplied to the synthesis gas steam mixture via line 220.

The combined synthesis gas and steam mixture is fed via line 222 to adistributor 224 disposed within a second sour shift vessel 226. Thedistributor is connected to a plurality of tubes 228 that passvertically through a bed of particulate Co/Mo sour shift catalyst 230.The synthesis gas steam mixture is able to pass from the distributorvertically through the tubes where it is heated thereby cooling thereactant gases in the catalyst bed 230. The tubes are connected to acollector 232 at the other end of the tubes that collects heatedsynthesis gas.

The heated synthesis is recovered from the vessel 226 via line 234 andmixed with a further amount of steam from line 236. The synthesis gassteam mixture is passed to heat exchanger 238 where its temperature isadjusted to the desired inlet temperature. The temperature adjustedsynthesis gas is fed from exchanger 238 via line 240 to a pre-shiftvessel 242 containing a first fixed bed of particulate sulphided Co/Mosulphur-tolerant water-gas shift catalyst 244. The flow of synthesis gascontaining steam is controlled such that the gas hourly space velocityin the first bed of catalyst is >12,500 h⁻¹. The synthesis gascontaining steam reacts over the catalyst to form carbon dioxide andhydrogen. The pre-shift vessel is operated adiabatically without coolingand the exothermic reactions heat the resulting pre-shifted gas stream.

The pre-shifted gas stream is recovered from the vessel 242 via line 246and passed through heat exchanger 248 where it is cooled.

The cooled pre-shifted gas stream is then fed via line 250 to the inletof the second water-gas shift vessel 226 containing the second fixed bedof particulate sulphided Co/Mo sulphur-tolerant water-gas shift catalyst230. The pre-shifted gas mixture is passed over the water-gas shiftcatalyst 230 further increasing the hydrogen content of the synthesisgas. The bed of catalyst 230 is cooled in heat exchange with thesynthesis gas/steam mixture 222 passing through the tubes 228 in adirection counter-current to the flow of pre-shifted gas stream throughthe vessel 226. A hydrogen-enriched shifted gas stream is recovered fromthe outlet of the second vessel 226 via line 252.

The shifted gas stream is cooled in heat exchangers 254 and 256 and thecooled shifted synthesis gas stream fed via line 258 to a thirdwater-gas shift vessel 260 containing a third particulate bed ofsulphur-tolerant Co/Mo water-gas shift catalyst 262. The shifted gasstream containing steam further reacts over the catalyst 262 to formcarbon dioxide and hydrogen. The third vessel is operated adiabaticallywithout cooling and the exothermic reactions heat the resulting shiftedgas stream. The shifted gas stream is recovered from the third water-gasshift vessel 260 and passed via line 264 to heat exchanger 266 where itis cooled. The cooled shifted gas stream is then fed via line 268 to afourth water-gas shift vessel 270 containing a fourth particulate bed ofsulphur-tolerant Co/Mo water-gas shift catalyst 272. The shifted gasstream containing steam further reacts over the catalyst 272 to formcarbon dioxide and hydrogen. The fourth vessel is operated adiabaticallywithout cooling and the exothermic reactions heat the resulting shiftedgas stream.

The shifted gas stream is recovered from the fourth water-gas shiftvessel 270 via line 274 and passed through heat exchanger 276, andoptionally further heat exchangers (not shown) to cool the gas below thedew point and so condense the remaining steam. The cooled shifted streamis fed via line 278 to separator 280 in which the condensate isseparated from the hydrogen rich shifted gas stream. The dryhydrogen-rich shifted gas stream is recovered from separator 280 vialine 282 and the condensate via line 284. The condensate may be used togenerate steam for use in the process. The dry hydrogen-rich shifted gasstream 282 may be used in downstream processing or sent to a gas washingunit (not shown) to recover CO2 and H2S and generate a hydrogen rich gasstream product. The carbon dioxide recovered from such processes may beused in carbon-capture and storage (CCS) processes or in enhanced oilrecovery (EOR) processes.

In an alternative embodiment, by utilising the collector 232 as thedistributor and vice-versa, the synthesis gas/steam mixture 222 may befed through the tubes 228 in a direction co-current with the flow ofpre-shifted gas through the second water-gas shift vessel 226.

The invention is further illustrated by reference to the followingcalculated Examples.

EXAMPLE 1 (COMPARATIVE)

The calculated mass balance below is for the use of three adiabatic sourshift reactors in series with cooling between reactors to carry out ahigh degree of shift (>90%) on a feed gas with a steam:CO ratio of 1.35and an R ratio of 0.37.

Reactor 1 Reactor 2 Reactor 3 Mol fraction In Out In Out In Out H2O0.38685 0.19883 0.19883 0.14571 0.14571 0.12987 CO 0.28383 0.095030.09503 0.04190 0.04190 0.02606 CO2 0.08589 0.27468 0.27468 0.327840.32784 0.34369 COS 0.00015 0.00006 0.00006 0.00002 0.00002 0.00001 H2S0.00545 0.00555 0.00555 0.00559 0.00559 0.00559 Argon 0.00583 0.005830.00583 0.00583 0.00583 0.00583 N2 0.00665 0.00665 0.00665 0.006650.00665 0.00665 NH3 0.00127 0.00127 0.00127 0.00127 0.00127 0.00127 CH40.00075 0.00114 0.00114 0.00116 0.00116 0.00116 H2 0.22333 0.410950.41095 0.46403 0.46403 0.47986 Flow kgmols/hr 31849.7 31825.0 31825.031823.0 31823.0 31823.0 T deg C. 238 439 220 277 220 237 P bar abs. 63.863 61.8 61.3 60.7 60.2

There is a large volume of catalyst in the first reactor, with a highexit temperature (˜440° C.), giving the potential for undesirable sidereactions, including methanation. This situation is exacerbated when thecatalyst is new and more active.

EXAMPLE 2

This is an example of the invention according to FIG. 1, based on thesame feed flow and composition and overall shift conversion duty asExample 1. The GHSV in the pre-shift vessel (128) is 41000 h⁻¹ and theGHSV in the gas-cooled converter (114) is 7500 h⁻¹.

Reactor 1 Reactor 2 Stream No 126 132 136 138 110 122 Mol fraction InOut In Cat Out Cat In tubes Out tubes H2O 0.38685 0.31650 0.316500.19037 0.38685 0.38685 CO 0.28383 0.21345 0.21345 0.08637 0.283830.28383 CO2 0.08589 0.15639 0.15639 0.28334 0.08589 0.08589 COS 0.000150.00002 0.00002 0.00004 0.00015 0.00015 H2S 0.00545 0.00559 0.005590.00557 0.00545 0.00545 Argon 0.00583 0.00583 0.00583 0.00583 0.005830.00583 N2 0.00665 0.00665 0.00665 0.00665 0.00665 0.00665 NH3 0.001270.00127 0.00127 0.00127 0.00127 0.00127 CH4 0.00075 0.00082 0.000820.00124 0.00075 0.00075 H2 0.22333 0.29350 0.29350 0.41932 0.223330.22333 Flow kgmols/hr 31849.7 31845.0 31845.0 31818.0 31849.7 31849.7 Tdeg C. 323 398 340 371.5 218 323 P bar abs. 63.8 62.8 62.5 62 63.9 63.8Reactor 3 Reactor 4 Stream No 144 150 154 160 Mol fraction In Out In OutH2O 0.19037 0.14567 0.14567 0.12990 CO 0.08637 0.04165 0.04165 0.02588CO2 0.28334 0.32807 0.32807 0.34385 COS 0.00004 0.00002 0.00002 0.00001H2S 0.00557 0.00559 0.00559 0.00559 Argon 0.00583 0.00583 0.005830.00583 N2 0.00665 0.00665 0.00665 0.00665 NH3 0.00127 0.00127 0.001270.00127 CH4 0.00124 0.00126 0.00126 0.00126 H2 0.41932 0.46399 0.463990.47975 Flow kgmols/hr 31818.0 31817.0 31817.0 31817.0 T deg C. 230 278220 237 P bar abs. 61.4 60.4 60.1 59.1

This process overcomes the problem of the high temperature seen inexample 1. The peak temperatures in shift stages 1 and 2 are about 400°C. and 375° C. respectively. Due to the fact that the temperatureprofile in stage 2 better follows that which is optimal for a highwater-gas shift reaction rate, the combined catalyst volume for reactors1 and 2 in this case is actually about 15% less than the volume ofcatalyst in reactor 1 in example 1, whereas the CO conversion isactually slightly higher.

EXAMPLE 3

This is an example of the invention according to FIG. 2. The syngas feed(steam:CO=0.21 and R=0.45) is preheated and some steam is added beforeit is fed to the tubes of the second shift reactor, where it flowscounter-current to the reactant gas flow, thereby cooling it. The heatedgas from the tubes is then cooled and mixed with further steam to give asteam:CO ratio of about 1.1. The synthesis gas containing steam is thenpassed to the first pre-shift reactor. The pre-shifted gas is cooledbefore passing it to the second (cooled) reactor. The shifted gas fromthe second shift stage flows to reactors 3 and 4 with cooling beforeeach stage, such that an overall CO conversion of about 92% is achieved.The peak temperature in each of the first two reactors is about 420° C.The GHSV in the pre-shift vessel (242) is 43000 h⁻¹ and the GHSV in thegas-cooled converter (226) is 6200 h⁻¹.

Feed Reactor 1 Reactor 2 Stream No 210 240 246 250 252 222 234 Molfraction In Out In Cat Out Cat In tubes Out tubes H2O 0.11546 0.394880.30714 0.30714 0.14146 0.25311 0.25311 CO 0.52476 0.35899 0.271660.27166 0.10579 0.44310 0.44310 CO2 0.01937 0.01325 0.10101 0.101010.26683 0.01636 0.01636 COS 0.00064 0.00044 0.00001 0.00001 0.000040.00054 0.00054 H2S 0.00744 0.00509 0.00551 0.00551 0.00548 0.006280.00628 Argon 0.00891 0.00609 0.00609 0.00609 0.00609 0.00752 0.00752 N20.05371 0.03674 0.03674 0.03674 0.03675 0.04535 0.04535 NH3 0.003030.00207 0.00207 0.00207 0.00207 0.00256 0.00256 CH4 0.00037 0.000250.00026 0.00026 0.00034 0.00031 0.00031 H2 0.26633 0.18219 0.293500.29350 0.43515 0.22488 0.22488 Flow kgmols/hr 19880 29060.0 29060.029060.0 29055.0 23544.0 23544.0 T deg C. 145 320 419 305 367 175 340 Pbar abs 38.9 38.6 37.8 37.5 36.5 38.9 38.8 Reactor 3 Reactor 4 Stream No258 264 268 274 Mol fraction In Out In Out H2O 0.14146 0.08205 0.082050.06569 CO 0.10579 0.04638 0.04638 0.03004 CO2 0.26683 0.32625 0.326250.34260 COS 0.00004 0.00003 0.00003 0.00002 H2S 0.00548 0.00550 0.005500.00551 Argon 0.00609 0.00609 0.00609 0.00609 N2 0.03675 0.03675 0.036750.03675 NH3 0.00207 0.00207 0.00207 0.00207 CH4 0.00034 0.00034 0.000340.00034 H2 0.43515 0.49455 0.49455 0.51089 Flow kgmols/hr 29055.029055.0 29055.0 29055.0 T deg C. 190 259 190 209 P bar abs. 35.9 34.934.6 33.7

EXAMPLE 4 (COMPARATIVE)

This is an example according to WO2010/013026 based on the same feedflow and composition, total steam addition and shift conversion duty asExample 3. The catalyst volume in the cooled first shift reactor isapproximately the same as the combined volume of the pre-shift andsecond shift reactor in Example 3. The synthesis gas feed (steam:CO=0.21and R=0.45) is preheated and some steam is added before it is fed to thetubes of the cooled shift reactor, where it flows co-current to thereactant gas flow, thereby cooling it. The heated gas from the tubes iscooled in an external heat exchanger and mixed with further steam togive a steam:CO ratio of about 1.1, before being fed to the catalyst.

In comparison to Example 3, there is the same CO conversion in thecooled reactor as the total conversion of the first two reactors inExample 3. However, the peak temperature (part way down the catalystbed) is about 440° C., which means a higher potential for side reactionsto occur.

Reactor 1 Mol fraction Syngas feed In Cat Out Cat In tubes Out tubes H2O0.115458 0.394882 0.141644 0.253105 0.253105 CO 0.524758 0.3589890.105823 0.443099 0.443099 CO2 0.019371 0.013252 0.26675 0.0163560.016356 COS 0.000643 0.00044 7.02E−05 0.000543 0.000543 H2S 0.0074360.005087 0.005458 0.006279 0.006279 Argon 0.008905 0.006092 0.0060940.007519 0.007519 N2 0.053706 0.03674 0.036751 0.045349 0.045349 NH30.003029 0.002072 0.002073 0.002558 0.002558 CH4 0.000367 0.0002510.000395 0.00031 0.00031 H2 0.266327 0.182195 0.434942 0.224883 0.224883Flow kgmols/hr 19880 29060.0 29052.0 23544.0 23544.0 T deg C. 145 305425 175 391 P bar abs. 38.9 37.6 36.5 38.9 38.8

EXAMPLE 5

This is an example of the invention according to FIG. 2, except thatthat shift reactor stages 3 and 4 are omitted in order to produce ashifted synthesis gas stream suited to methanol synthesis.

The syngas feed (steam:CO=0.15 and R=0.45) is split into a main flow andbypass flow. The main flow (91%) is preheated and is fed to the tubes ofthe second shift reactor, where it flows co-current to the reactant gasflow, thereby cooling it. The heated gas from the tubes is then cooledand mixed with steam to give a steam:CO ratio of about 0.82 and thenpasses to the first (pre-shift reactor). The pre-shifted gas is cooledbefore passing to the second (cooled) reactor. The overall conversion ofCO in the gas passing through the two shift reactors is about 60%. Theshifted gas leaving the second stage is cooled and mixed with theun-shifted by-pass syngas feed stream.

This combined stream is further cooled to condense water, which isremoved as condensate, before being fed to an Acid Gas Removal unit toremove CO2 and sulphur compounds to make a gas with a suitablecomposition to utilise for methanol synthesis.

Syngas feed Reactor 1 Reactor 2 Stream No 210 240 246 250 252 222 234Mol fraction In Cat Out Cat In Cat Out Cat In tubes Out tubes H2O0.082934 0.326045 0.21593 0.21593 0.08453 0.082934 0.082934 CO 0.5440640.399835 0.29016 0.29016 0.15881 0.544064 0.544064 CO2 0.020081 0.0147570.12490 0.12490 0.25630 0.020081 0.020081 COS 0.000666 0.000489 0.000020.00002 0.00011 0.000666 0.000666 H2S 0.007703 0.005661 0.00613 0.006130.00604 0.007703 0.007703 Argon 0.009232 0.006785 0.00679 0.006790.00679 0.009232 0.009232 N2 0.055682 0.040921 0.04092 0.04092 0.040950.055682 0.055682 NH3 0.003133 0.002302 0.00230 0.00230 0.00230 0.0031330.003133 CH4 0.000381 0.00028 0.00029 0.00029 0.00047 0.000381 0.000381H2 0.276124 0.202925 0.31257 0.31257 0.44369 0.276124 0.276124 Flowkgmols/hr 19174 23741.0 23741.0 23741.0 23741.0 17447 17447 T deg C. 131290 418 390 412 200 392 P bar abs. 38.9 38.3 38 37.8 37.3 38.6 38.5

1-18. (canceled)
 19. A process for increasing the hydrogen content of asynthesis gas containing one or more sulphur compounds, said synthesisgas comprising hydrogen, carbon oxides and steam, and having a ratio, R,defined as R═(H₂—CO₂)/(CO+CO₂)≦0.6 and a steam to carbon monoxide ratio1.8, comprising the steps of (i) heating the synthesis gas, (ii) passingat least a portion of the heated synthesis gas adiabatically through afirst bed of sulphur-tolerant water-gas shift catalyst disposed in afirst shift vessel at a gas hourly space velocity ≧12,500 hour⁻¹to forma pre-shifted gas stream, and (iii) forming a shifted gas stream bysubjecting at least a portion of the pre-shifted gas stream to a secondstage of water-gas shift in a second shift vessel containing a secondbed of sulphur-tolerant water-gas shift catalyst that is cooled in heatexchange with a gas stream comprising the synthesis gas.
 20. A processaccording to claim 19 wherein the synthesis gas containing one or moresulphur compounds is formed by gasification of a carbonaceous feedstockat elevated temperature and pressure, followed by cooling and washingthe resulting gas stream to remove particulate material.
 21. A processaccording to claim 20 wherein the gasification is performed on a coalpowder or aqueous slurry in a gasifier using oxygen or air and in thepresence of steam at a pressure up to about 85 bar abs and an exittemperature up to 1450° C.
 22. A process according to claim 20 whereinthe steam to carbon monoxide ratio is in the range 0.2 to 1.8.
 23. Aprocess according to claim 20 wherein the R ratio is in the range 0.1 to0.6.
 24. A process according to claim 19 wherein the space velocity ofthe synthesis gas flowing through the first bed of sulphur-tolerantwater-gas shift catalyst is ≧15,500 hour⁻¹.
 25. A process according toclaim 19 wherein the inlet temperature for the first bed of water-gasshift catalyst is in the range 190 to 350° C.
 26. A process according toclaim 19 wherein the synthesis gas is heated by passing it through aplurality of tubes, coils or plates disposed within the second catalystbed.
 27. A process according to claim 19 wherein the synthesis gas issubjected to the water-gas shift reaction over a supportedcobalt-molybdenum water-gas shift catalyst.
 28. A process according toclaim 19 wherein the water-gas shift reaction over the first bed ofwater-gas shift catalyst is performed at a temperature in the range 190to 420° C.
 29. A process according to claim 19 wherein the heatedsynthesis gas is divided into first and second streams, with the firststream passed over the first bed of shift catalyst and the second streamby-passing the first bed of shift catalyst, thereby forming a catalystby-pass stream.
 30. A process according to claim 29 wherein the by-passstream is fed to one or more of the pre-shifted gas stream, the shiftedgas stream, or separately to downstream processes.
 31. A processaccording to claim 19 wherein the second stage of water-gas shift isperformed in a vessel containing a supported cobalt-molybdenum water-gasshift catalyst.
 32. A process according to claim 19 wherein the secondstage of water gas shift is performed at a temperature in the range 250to 420° C.
 33. A process according to claim 19 wherein the spacevelocity in the second bed of sulphur-tolerant water-gas shift catalystis ≧6000 h⁻¹.
 34. A process according to claim 19 wherein the synthesisgas is passed through the tubes in the second shift vessel in adirection that is counter-current or co-current to the flow of thepre-shifted gas stream through the vessel.
 35. A process according toclaim 19 wherein the shifted gas stream is subjected to one or morefurther water-gas shift stages.
 36. A process according to claim 19further comprising the steps of: (i) cooling a shifted gas streamobtained from the one or more further stages of water-gas shift, or amixture of the shifted gas stream and a bypass stream, to below the dewpoint to condense water, (ii) separating the resulting condensatetherefrom to form a dry shifted gas stream, (iii) feeding the dryshifted gas stream to a gas-washing unit operating by means ofcounter-current solvent flow, to produce a product synthesis gasenriched in hydrogen and (iv) collecting the product synthesis gas fromthe washing unit.
 37. A process according to claim 20 wherein the Rratio is in the range 0.2 to 0.6.
 38. A process according to claim 19wherein the space velocity of the synthesis gas flowing through thefirst bed of sulphur-tolerant water-gas shift catalyst is ≧17,500hour⁻¹.
 39. A process according to claim 19 wherein the space velocityof the synthesis gas flowing through the first bed of sulphur-tolerantwater-gas shift catalyst is ≧20,000 hour⁻¹.
 40. A process according toclaim 19 wherein the space velocity in the second bed ofsulphur-tolerant water-gas shift catalyst is in the range 6000 h⁻¹ to12000 h⁻¹.